Hydrogen Management for Hydroprocessing Units

ABSTRACT

Improved hydroprocessing processes for upgrading refinery streams via the use of rapid cycle pressure swing absorption having a cycle time of less than 30 s for increasing the concentration of hydrogen in the vapor phase product recycled to the hydroprocessing zone.

FIELD OF THE INVENTION

This invention relates to improved hydroprocessing processes forupgrading refinery streams via the use of rapid cycle pressure swingadsorption having a cycle time of less than one minute for increasingthe concentration of hydrogen for use in hydroprocessing units.

BACKGROUND OF THE INVENTION

Hydroprocessing processes are used by petroleum refiners to improve theproperties and hence value of many refinery streams. Suchhydroprocessing units include hydrotreating, hydrocracking,hydroisomerization and hydrogenation process units. Hydroprocessing isgenerally accomplished by contacting a hydrocarbon feedstock in ahydroprocessing reaction vessel, or zone, with a suitablehydroprocessing catalyst under hydroprocessing conditions of elevatedtemperature and pressure in the presence of a hydrogen-containing treatgas to yield an upgraded product having the desired product properties,such as sulfur and nitrogen levels, boiling point, aromaticconcentration, pour point and viscosity index. The operating conditionsand the hydroprocessing catalysts used will influence the quality of theresulting hydroprocessing products.

Several types of hydroprocessing operations are practiced commerciallyin refining operations. For example, hydrotreating is typically used toremove heteroatoms, such as sulfur and nitrogen, from hydrocarbonfeedstreams such as naphtha, kerosene, diesel, gas oil, vacuum gas oil(VGO), and residua, by contacting the feedstream with hydrogen and asuitable hydrotreating catalyst, at hydrotreating conditions oftemperature, pressure and flow rates to result in the heteroatoms beingconverted to hydrogen sulfide. Hydrotreaters are also employed toimprove other properties of hydrocarbon streams in the refinery.Hydrocracking is typically used to remove sulfur and nitrogen, and toreduce the boiling point of heavier molecules by converting them intolighter molecules, by contacting the feedstream with hydrogen and asuitable hydrocracking catalyst, at hydrocracking process conditions.Hydrodewaxing and hydroisomerization of distillate and lubricating oilsmodifies the molecular structure and hence the pour point of thesemolecules, by contacting the feedstream with hydrogen over a suitablecatalyst, at hydrodewaxing and hydroisomerization process conditions.Hydroprocessing for olefin and aromatic saturation reduces theconcentration of aromatics and olefins by contacting the feedstream withhydrogen over a suitable catalyst at aromatic/olefin saturationconditions.

All of these hydroprocessing operations require the use of hydrogen, andthe amount of hydrogen required to operate these hydroprocessing unitshas greatly increased for several reasons. Regulatory pressure in theUnited States, Europe, Asia, and elsewhere has resulted in a trend toincreasingly severe and/or selective hydroprocessing processes to formhydrocarbon products having very low levels of sulfur and other tailoredproperties, such as reduced aromatics levels, and improved pour pointand viscosity index. The move to process heavier crude oils and thereduced market for fuel oil is increasing the need for hydrocracking,again leading to a higher hydrogen demand. As the qualities oflubricating oils improve, the need to remove even more sulfur, reducearomatics levels, and improve pour point and viscosity index haveincreased the need for hydroprocessing. Further, many refineries receivelarge amounts of hydrogen as a by-product of catalytic reforming ontheir site. However, current treads to reduce aromatics in gasoline areconstraining the use of catalytic reforming and thus removing a sourceof hydrogen. Thus, there is an ever growing need for improved hydrogenmanagement associated with the various process units.

Hydroprocessing units use relatively large quantities of hydrogen thatare often obtained from process units that generate hydrogen, either asa main product stream or as a side product stream. The vapor phaseproduct stream from hydroprocessing units typically contains unreactedhydrogen that is recycled to the hydroprocessing reaction zone. Sincehydrogen is an important reactant in hydroprocessing, economic means topurify hydrogen in hydrogen-containing streams used as feed streamsand/or as recycle streams is desirable. A greater concentration ofhydrogen in either of these two types of hydrogen-containing streamsallows for a more efficient process with higher feed throughput.

The type of feed to be processed, product quality requirements, yield,and the amount of conversion for a specific catalyst cycle lifedetermines the hydrogen partial pressure required for the operation of ahydroprocessing unit. The unit's operating pressure and the recycle gaspurity determine the hydrogen partial pressure of the hydroprocessingunit. Since there is limited control over the composition of the flashedgas from the downstream hydroprocessor separator or flash drum, thehydrogen composition of the recycle flash gas limits the hydrogenpartial pressure ultimately delivered to the hydroprocessor reactor. Arelatively lower hydrogen partial pressure in the recycle gas streameffectively lowers the partial pressure of the hydrogen gas inputcomponent to the reactor and thereby adversely affects the operatingperformance with respect to product quantity and quality, catalyst cyclelife, etc. To offset this lower performance, the operating pressure ofthe hydroprocessor reactor has to be increased, which can be undesirablefrom an operational point of view. Conversely, by increasing theefficiency of hydrogen gas recovery and hydrogen concentration, thehydrogen partial pressure of the recycle gas stream is improved. Thisresults in an overall improved performance of the hydroprocessingprocess unit as measured by these parameters.

Some conventional methods have been proposed that attempt to improve thehydrogen utilization efficiency of the hydroprocessing unit byincreasing the concentration of the hydrogen in the recycle gas stream.Such processes typically result in significant additional equipmentcosts and/or require significant changes in operating conditions, suchas temperature and pressure, which typically results in increasedcapital and operating costs.

One process that has been adopted to improve the hydrogen purity of therecycle stream in a hydroprocessing unit is conventional pressure swingadsorption (CPSA). See, for example, U.S. Pat. No. 4,457,384 issued Jul.3, 1984 to Lummus Crest, Inc. However, in order to incorporate the PSAunit, the pressure of the reactor effluent gas stream must be reducedfrom about 2,500 psig (175.8 kg/cm²) to about 350 psig (24.6 kg/cm²).Although the purity of the recycle hydrogen stream can be increased toabout 99 mol %, the recycled gaseous stream must be subjected tocompression to return it to 2,500 psig (175.8 kg/cm²) beforeintroduction into the hydroprocessing feed stream. The net result isthat the capital, operating and maintenance costs are substantiallyincreased by the addition of a large compressor that is required whenusing a conventional PSA unit.

Another method is described in U.S. Pat. No. 4,362,613 to MacLean whichis incorporated herein by reference. MacLean uses membranes withpressure drops up to 150 atmospheres and which also incurs substantialcapital investment and operating costs.

There is therefore a need for an improved process for enhancing theefficiency of hydrogen utilization by means that are compatible withexisting hydroprocessing units. It is desired that such a process wouldnot adversely affect the hydroprocessor throughput or the overalleconomies of the system, including capital expenditures and operatingexpenditures, the latter including maintenance and energy consumption.

In other words, although various hydroprocessing processes are practicedcommercially, there is still a need in the art for improvedhydroprocessing processes that can be practiced more efficiently andwith higher reactor throughput by combining improvements to hydrogenrecovery and purification with hydroprocessing units.

SUMMARY OF THE INVENTION

In a preferred embodiment, there is provided a process for upgrading ahydrocarbon feed in a hydroprocessing process unit, comprising:

a) contacting said hydrocarbon feed in a hydroprocessing zone withhydrogen, a portion of which is obtained from a hydrogen-containingmake-up gas, and a catalytically effective amount of a hydroprocessingcatalyst at hydroprocessing conditions thereby resulting in a liquidphase and a vapor phase product;

b) separating said liquid phase and said vapor phase, which vapor phasecontains hydrogen and light hydrocarbons;

c) removing at least a portion of the light hydrocarbons from thehydrogen-containing make-up treat gas, the vapor phase product, or both,in a rapid cycle pressure swing adsorption unit containing a pluralityof adsorbent beds and having a total cycle time of less than about 30seconds and a pressure drop within each adsorbent bed of greater thanabout 5 inches of water per foot of bed length;

d) recycling at least a portion of the vapor phase of step c) abovehaving a higher concentration of hydrogen to the hydroprocessing zone.

In another preferred embodiment, the hydrocarbon feed is selected fromthe group consisting of naphtha boiling range feeds, kerosene and jetfuel boiling range feeds, distillate boiling range feeds, resides andcrudes.

In yet another preferred embodiment, the total cycle time or the rapidcycle pressure swing adsorption step is less than about 15 seconds.

In still another preferred embodiment the total cycle time is less thanabout 10 seconds and the pressure drop is greater than about 10 inchesof water per foot of bed length for the rapid cycle pressure swingadsorption step.

DETAILED DESCRIPTION OF THE INVENTION

It has been recognized that by increasing the efficient use of hydrogen,existing equipment could be employed to increase the throughput of thefeedstock resulting in higher product yields. A further advantage to themore efficient utilization of hydrogen is the reduction in the amount ofmake-up hydrogen that must be provided by, for example, a hydrogenplant, cryo-unit or reformer.

The instant invention is applicable to any unit in a petroleum refinerythat uses hydrogen as a treat-gas stream, or as a recycle stream, orproduces hydrogen as a primary product or as a side product stream. Itis particularly applicable to those process units that use hydrogen as areactant to upgrade or to convert a hydrocarbon stream to lower boilingproducts. Such process units are typically referred to ashydroprocessing units. The art has long recognized the importance ofimproving the purity (concentration) of hydrogen in the recycle streamof hydroprocessing units. Non-limiting types of hydroprocessing that areincluded herein are: hydrotreating wherein light hydrocarbon, naphtha,diesel, distillate, atmospheric and vacuum gas oils, kerosene, jet,cycle oils, lubestock and waxes, atmospheric and vacuum residua,pyrolysis gasoline, and crude streams are upgraded by the removal ofheteroatoms, hydrogenation wherein double bonds are converted to olefinsand paraffins and aromatics are saturated to naphthenes as well as theremoval of at least a portion of heteroatoms, hydrocracking wherein highboiling streams are converted to more valuable lower boiling streams,hydroisomerization wherein paraffinic compounds are converted toisoparaffins, hydrofinishing, which is a mild hydrotreating process usedparticularly to replace or supplement clay treating of lube oils andwaxes. Other hydroprocessing process that are incorporated within thisinvention include catalytic dewaxing which is a catalytic hydrocrackingprocess that uses molecular sieve catalysts to selectively hydrocrackwaxes present in a feedstock into lighter hydrocarbon fractions; waxhydroisomerization wherein wax molecules are converted to branchedmolecules in a catalytic reaction and converted into high VI lubricants.

Also, lubricating and/or specialty oil stocks such as deasphalted oilstocks, lube oil distillates, and solvent extracted lube oil raffinatescan have their viscosity indexes increased by hydrotreating, employingspecific bulk metal sulfide hydrotreating catalysts selected from thegroup consisting of bulk Cr/Ni/Mo sulfide catalyst, bulk Ni/Mo/Mnsulfide catalyst and mixtures thereof wherein the catalysts are preparedfrom specific metal complexes and wherein the Ni/Mn/Mo sulfide catalystis prepared from the oxide precursor decomposed in an inert atmospheresuch as N₂ and subsequently sulfided using H₂S/H₂ and the Cr/Ni/Mosulfide catalyst is prepared from the sulfide precursor and decomposedin a non-oxidizing, sulfur containing atmosphere.

Herein, the term “hydrocarbon feed” is defined as a refinery, chemicalor other industrial plant stream that is comprised of hydrocarbonsincluding such streams wherein small levels (less than 5%) ofnon-hydrocarbon contaminants such as, but not limited to, sulfur, water,ammonia, and metals may be present in the hydrocarbon feed. The term“light hydrocarbons” means a hydrocarbon mixture comprised ofhydrocarbon compounds of about 1 to about 5 carbon atoms in weight(i.e., C₁ to C₅ weight hydrocarbon compounds). It will be understoodthat the terms “hydrocarbon” and “hydrocarbonaceous” are usedinterchangeably herein when referring to feedstreams.

Feedstreams that can be hydroprocessed in accordance with the presentinvention are any hydrocarbonaceous feedstreams that are upgraded byhydroprocessing. Non-limiting examples of such feedstreams include lighthydrocarbon boiling range feedstreams, naphtha boiling rangefeedstreams, kerosene and jet boiling range feedstreams, diesel anddistillate boiling range feedstreams, cycle oils produced from the FluidCatalytic Cracker (FCC), atmospheric and vacuum gas oils, atmosphericand vacuum residua, pyrolysis gasoline, Fischer-Tropsch liquids,raffinates, waxes, lube oils, and crudes, as well as heavier gas oil andresid boiling range feedstreams.

For example, in the case of hydrotreating, heteroatoms such as sulfurand nitrogen are typically removed from the aforementioned feed streams,whereas in the case of hydrocracking heavier boiling range gas oil andreside type streams are converted to lower boiling product streams.Non-limiting examples of naphtha feedstreams that can be treated inaccordance with the present invention are those containing componentsboiling in the range from about 50° F. to about 450° F., at atmosphericpressure. The naphtha feedstream generally contains cracked naphthawhich usually comprises fluid catalytic cracking unit naphtha (FCCcatalytic naphtha), coker naphtha, hydrocracker naphtha, residhydrotreater naphtha, debutanized natural gasoline (DNG), and gasolineblending components from other sources wherein a naphtha boiling rangestream can be produced. Non-limiting examples of distillate feedstreamsthat can be treated in accordance with the present invention are thoseboiling in the range of about 288° C. (550° F.), such as atmospheric gasoils, vacuum gas oils, deasphalted vacuum and atmospheric residua,mildly cracked residual oils, coker distillates, straight rundistillates, solvent-deasphalted oils, pyrolysis-derived oils, highboiling synthetic oils, cycle oils and cat cracker distillates. Apreferred hydrotreating feedstock is a gas oil or other hydrocarbonfraction having at least 50% by weight, and most usually at least 75% byweight of its components boiling at temperatures between about 316° C.(600° F.) and 538° C. (1000° F.). Crude oils can also be feed inaccordance with the present invention.

Illustrative hydrocarbon feedstreams that are upgraded by hydrocrackinginclude those containing components boiling above about 260° C. (500°F.), such as Fischer-Tropsch liquids, atmospheric gas oils, vacuum gasoils, deasphalted, vacuum, and atmospheric residua, hydrotreated ormildly hydrocracked residual oils, coker distillates, straight rundistillates, solvent-deasphalted oils, pyrolysis-derived oils, highboiling synthetic oils, cycle oils and cat cracker distillates. Apreferred hydrocracking feedstream is a gas oil or other hydrocarbonfraction having at least 50% by weight, and most usually at least 75% byweight, of its components boiling at temperatures above the end point ofthe desired product. One of the most preferred gas oil feedstreams willcontain hydrocarbon components that boil above 260° C. (500° F.), withbest results being achieved with feeds containing at least 25 percent byvolume of the components boiling between about 315° C. (600° F.) and538° C. (1000° F.). A preferred heavy feedstream boils in the range fromabout 93° C. to about 565° C. (200-1050° F.). Hydroisomerizationfeedstreams are typically paraffinic, such as wax streams, particularlyFischer-Tropsch waxes and light paraffins.

The term “hydrotreating” as used herein refers to processes wherein ahydrogen-containing treat gas is used in the presence of suitablecatalysts which are primarily active for the removal of heteroatoms,such as sulfur and nitrogen and for some hydrogenation of aromatics.Suitable hydrotreating catalysts for use in the present invention areany known conventional hydrotreating catalysts and include those whichare comprised of at least one Group VIII metal, preferably iron, cobaltand nickel, more preferably cobalt and/or nickel and at least one GroupVI metal, preferably molybdenum and tungsten, either as a bulk catalyst,or supported on a high surface area support material, preferablyalumina. Other suitable hydrotreating catalysts include zeoliticcatalysts, as well as noble metal catalysts where the noble metal isselected from palladium and platinum. It is within the scope of thepresent invention that more than one type of hydrotreating catalyst beused in the same reaction vessel. The Group VIII metal is typicallypresent in an amount ranging from about 2 to about 20 wt. %, preferablyfrom about 4 to about 12 wt. %. The Group VI metal will typically bepresent in an amount ranging from about 1 to about 25 wt-%, preferablyfrom about 2 to about 25 wt. %. As previously mentioned, typicalhydrotreating temperatures range from about 204° C. (400° F.) to about482° C. (900° F.) with pressures from about 3.5 MPa (500 psig) to about17.3 MPa (2500 psig), preferably from about 3.5 MPa (500 psig) to about13.8 MPa (2000 psig) and a liquid hourly space velocity of thefeedstream from about 0.1 hr⁻¹ to about 10 hr⁻¹.

The active metals employed in the preferred hydrocracking catalysts ofthe present invention as hydrogenation components are those of GroupVIII of the Periodic Table of the Elements, i.e., iron, cobalt, nickel,ruthenium, rhodium, palladium, osmium, iridium and platinum. One or morepromoter metals can also be present. Preferred promoter metals are thosefrom Group VIB, e.g., molybdenum and tungsten, more preferablymolybdenum. The amount of hydrogenation metal component in the catalystcan vary within wide ranges. Broadly speaking, any amount between about0.05 percent and 30 percent by weight may be used. In the case of thenoble metals, it is preferred to use about 0.05 to about 2 weightpercent of such metals. The preferred method for incorporating thehydrogenation metal component is to contact a zeolite base material,preferably a zeolite with the Faujasite or Beta zeolite structure, withan aqueous solution of a suitable compound of the desired metal whereinthe metal is present in a cationic form. Following addition of theselected hydrogenation metal or metals, the resulting catalyst powder isthen filtered, dried, pelleted with added lubricants, binders or thelike if desired, and calcined in air at temperatures of, e.g., 371°-648°C. (700°-1200° F.) in order to activate the catalyst and decomposeammonium ions. Alternatively, the zeolite component may first bepelleted, followed by the addition of the hydrogenating component andactivation by calcining. The foregoing catalysts may be employed inundiluted form, or the powdered zeolite catalyst may be mixed andcopelleted with other relatively less active catalysts, diluents orbinders such as alumina, silica gel, silica-alumina cogels, activatedclays and the like in proportions ranging between 5 and 90 weightpercent. These diluents may be employed as such or they may contain aminor proportion of an added hydrogenating metal such as a Group VIBand/or Group VIII metal. Additional metal promoted hydrocrackingcatalysts may also be utilized in the process of the present inventionwhich comprises, for example, aluminophosphate molecular sieves,crystalline chromosilicates and other crystalline silicates. Crystallinechromosilicates are more fully described in U.S. Pat. No. 4,363,718(Klotz).

Hydrocracking is typically performed at a temperature from about 232° C.(450° F.) to about 468° C. (875° F.), at a pressure from about 3.6 MPa(500 psig) to about 20.8 MPa (3000 psig), at a liquid hourly spacevelocity (LHSV) from about 0.1 to about 30 hr⁻¹, and at a hydrogencirculation rate from about 337 normal m³/m³ (2000 standard cubic feetper barrel) to about 4200 normal m³/m³ (25,000 standard cubic feet perbarrel). In accordance with the present invention, the term “substantialconversion to lower boiling products” is meant to connote the conversionof at least 5 volume percent of the fresh feedstock to lower boilingproducts. In a preferred embodiment, the per pass conversion in thehydrocracking zone is in the range from about 15% to about 45%. Morepreferably the per pass conversion is in the range from about 20% toabout 40%.

The hydroisomerization of the hydrocarbon feedstock is performed in ahydroisomerization zone which includes a hydroisomerization catalyst,the presence of hydrogen, and which is operated under hydroisomerizationconditions sufficient to hydrogenate diolefins to mono-olefins and toisomerize mono-olefins. Preferably, the hydroisomerization conditionsinclude a temperature in the range of from about 0° F. to about 500° F.,more preferably from about 75° F. to about 400° F., and most preferablyfrom 100° F. to 200° F.; a pressure in the range of from about 100 psigto about 1500 psig, more preferably from about 150 psig to about 1000psig, and most preferably from 200 psig to 600 psig; and a liquid hourlyspace velocity (LHSV) in the range of from about 0.01 hr⁻¹ to about 100hr⁻¹, more preferably from 1 hr⁻¹ to about 50 hr⁻¹, and most preferablyfrom 5 hr⁻¹ to 15 hr⁻¹. Hydroisomerization of paraffinic hydrocarbonstypically employs a catalyst composed of a noble metal, alumina andchlorine, said catalyst prepared by treating a composite of a noblemetal and alumina with an inorganic or organic salt of aluminum,preferably aluminum nitrate, calcining the treated composite andthereafter contacting the composite with a conventional chlorideactivating agent.

Wax hydroisomerization is also an important process, especially whenconverting slack waxes as well as Fischer-Tropsch waxes to more valuablefuel and lube products have acceptable pour points with a high viscosityindex. Waxes are typically hydroisomerized using a catalyst containing ahydrogenating metal component-typically one from Group IV, or Group VIIIof the Periodic Table, or mixtures thereof. The reaction is conductedunder conditions of temperature between about 500° F. to 750° F.,preferably between about 570° F. to 680° F., and pressures of from about500 to 3000 psi H₂ preferably from about 500-1500 psi H₂, at hydrogengas rates from 1000 to 10,000 SCF/bbl, and at space velocities in therange of from 0.1 to 10 v/v/hr, preferably from 0.5 to 2 v/v/hr.Following hydroisomerization, the isomerate is fractionated into a lubescut and a fuels cut. The lubes cut can then be dewaxed to recoverunconverted wax.

In Conventional Pressure Swing Adsorption (“conventional PSA”) a gaseousmixture is conducted under pressure for a period of time over a firstbed of a solid sorbent that is selective or relatively selective for oneor more components, usually regarded as a contaminant that is to beremoved from the gas stream. It is possible to remove two or morecontaminants simultaneously but for convenience, the component orcomponents that are to be removed will be referred to in the singularand referred to as a contaminant. The gaseous mixture is passed over afirst adsorption bed in a first vessel and emerges from the bed depletedin the contaminant that remains sorbed in the bed. After a predeterminedtime or, alternatively when a break-through of the contaminant isobserved, the flow of the gaseous mixture is switched to a secondadsorption bed in a second vessel for the purification to continue.While the second bed is in adsorption service, the sorbed contaminant isremoved from the first adsorption bed by a reduction in pressure,usually accompanied by a reverse flow of gas to desorb the contaminant.As the pressure in the vessels is reduced, the contaminant previouslyadsorbed on the bed is progressively desorbed into the tail gas systemthat typically comprises a large tail gas drum, together with a controlsystem designed to minimize pressure fluctuations to downstream systems.The contaminant can be collected from the tail gas system in anysuitable manner and processed further or disposed of as appropriate.When desorption is complete, the sorbent bed may be purged with an inertgas stream, e.g., nitrogen or a purified stream of the process gas.Purging may be facilitated by the use of a higher temperature purge gasstream.

After, e.g., breakthrough in the second bed, and after the first bed hasbeen regenerated so that it is again prepared for adsorption service,the flow of the gaseous mixture is switched from the second bed to thefirst bed, and the second bed is regenerated. The total cycle time isthe length of time from when the gaseous mixture is first conducted tothe first bed in a first cycle to the time when the gaseous mixture isfirst conducted to the first bed in the immediately succeeding cycle,i.e., after a single regeneration of the first bed. The use of third,fourth, fifth, etc. vessels in addition to the second vessel, as mightbe needed when adsorption time is short but desorption time is long,will serve to increase cycle time.

Thus, in one configuration, a pressure swing cycle will include a feedstep, at least one depressurization step, a purge step, and finally arepressurization step to prepare the adsorbent material forreintroduction of the feed step. The sorption of the contaminantsusually takes place by physical sorption onto the sorbent that isnormally a porous solid such as alumina, silica or silica-alumina thathas an affinity for the contaminant. Zeolites are often used in manyapplications since they may exhibit a significant degree of selectivityfor certain contaminants by reason of their controlled and predictablepore sizes. Normally, chemical reaction with the sorbent is not favoredin view of the increased difficulty of achieving desorption of specieswhich have become chemically bound to the sorbent but chemisorption ismy no means to be excluded if the sorbed materials may be effectivelydesorbed during the desorption portion of the cycle, e.g., by the use ofhigher temperatures coupled with the reduction in pressure.

Conventional PSA is not suitable for use in the present invention for avariety of reasons. For example, conventional PSA units are costly tobuild and operate and are much large in size for the amount of hydrogenthat needs to be recovered from such streams as compared to RCPSA. Also,a conventional pressure swing adsorption unit will generally have cycletimes in excess of one minute, typically in excess of 2 to 4 minutes dueto time limitations required to allow diffusion of the componentsthrough the larger beds utilized in conventional PSA and the equipmentconfiguration and valving involved. Instead, rapid cycle pressure swingadsorption is utilized which has cycle times of less than one minute.The total cycle times may be less than 30 seconds, preferably less than15 seconds, more preferably less than 10 seconds, even more preferablyless than 5 seconds, and even more preferably less 2 seconds. Further,the rapid cycle pressure swing adsorption units used can make use ofsubstantially different sorbents, such as, but not limited to,structured materials such as monoliths.

The overall adsorption rate of the adsorption processes, whetherconventional PSA or RCPSA, is characterized by the mass transfer rateconstant in the gas phase (τ_(g)) and the mass transfer rate constant inthe solid phase (τ_(s)). A material's mass transfer rates of a materialare dependent upon the adsorbent, the adsorbed compound, the pressureand the temperature. The mass transfer rate constant in the gas phase isdefined as:

τ_(g) =D _(g) /R _(g) ² (in cm²/sec)  (1)

where D_(g) is the diffusion coefficient in the gas phase and R_(g) isthe characteristic dimension of the gas medium. Here the gas diffusionin the gas phase, D_(g), is well known in the art and the characteristicdimension of the gas medium, R_(g) is defined as the channel widthbetween two layers of the structured adsorbent material.

The mass transfer rate constant in the solid phase of a material isdefined as:

τ_(s) =D _(s) /R _(s) ² (in cm²/sec)  (2)

where D_(s) is the diffusion coefficient in the solid phase and R_(s) isthe characteristic dimension of the solid medium. Here the gas diffusioncoefficient in the solid phase, D_(s), is well known in the art and thecharacteristic dimension of the solid medium, R_(s) is defined as thewidth of the adsorbent layer.

D. M. Ruthven & C. Thaeron, Performance of a Parallel Passage AbsorbentContactor, Separation and Purification Technology 12 (1997) 43-60, whichis incorporated by reference, clarifies that for flow through a monolithor a structured adsorbent that channel width is a good characteristicdimension for the gas medium, R_(g). U.S. Pat. No. 6,607,584 to Moreauet al., which is incorporated by reference, also describes the detailsfor calculating these transfer rates and associated coefficients for agiven adsorbent and standard stream composition for conventional PSA.Calculation of these mass transfer rate constants is well known to oneof ordinary skill in the art and may also be derived by one of ordinaryskill in the art from standard testing data.

Conventional PSA relies on the use of adsorbent beds of particulateadsorbents. Additionally, due to construction constraints, conventionalPSA is usually comprised of 2 or more separate beds that cycle so thatat least one or more beds is fully or at least partially in the feedportion of the cycle at any one time in order to limit disruptions orsurges in the treated process flow. However, due to the relatively largesize of conventional PSA equipment, the particle size of the adsorbentmaterial is general limited particle sizes of about 1 mm and above.Otherwise, excessive pressure drop, increased cycle times, limiteddesorption, and channeling of feed materials will result.

RCPSA utilizes a rotary valving system to conduct the gas flow through arotary sorber module that contains a number of separate compartmentseach of which is successively cycled through the sorption and desorptionsteps as the rotary module completes the cycle of operations. The rotarysorber module is normally comprised of tubes held between two sealplates on either end of the rotary sorber module wherein the seal platesare in contact with a stator comprised of separate manifolds wherein theinlet gas is conducted to the RCPSA tubes and processed purified productgas and the tail gas exiting the RCPSA tubes is conducted away fromrotary sorber module. By suitable arrangement of the seal plates andmanifolds, a number of individual compartments may be passing throughthe characteristic steps of the complete cycle at any one time. Incontrast with conventional PSA, the flow and pressure variationsrequired for the sorption/desorption cycle may be changed in a number ofseparate increments on the order of seconds per cycle, which smoothesout the pressure and flow rate pulsations encountered by the compressionand valving machinery. In this form, the RCPSA module includes valvingelements angularly spaced around the circular path taken by the rotatingsorption module so that each compartment is successively passed to a gasflow path in the appropriate direction and pressure to achieve one ofthe incremental pressure/flow direction steps in the complete RCPSAcycle. A key advantage of the RCPSA technology is a much more efficientuse of the adsorbent material. Since the quantity of adsorbent requiredwith RCPSA technology can be only a fraction of that required forconventional PSA technology to achieve the same separation quantitiesand qualities. The footprint, investment, and the amount of activeadsorbent required for RCPSA is significantly lower than that for aconventional PSA unit processing an equivalent amount of gas.

In an embodiment, RCPSA bed length unit pressure drops, requiredadsorption activities, and mechanical constraints (due to centrifugalacceleration of the rotating beds in RCPSA), prevent the use of manyconventional PSA adsorbent bed materials, in particular adsorbents thatare in a loose pelletized, particulate, beaded, or extrudate form. In apreferred embodiment, adsorbent materials are secured to a supportingunderstructure material for use in an RCPSA rotating apparatus. Forexample, one embodiment of the rotary RCPSA apparatus can be in the formof adsorbent sheets comprising adsorbent material coupled to astructured reinforcement material. A suitable binder may be used toattach the adsorbent material to the reinforcement material.Non-limiting examples of reinforcement material include monoliths, amineral fiber matrix, (such as a glass fiber matrix), a metal wirematrix (such as a wire mesh screen), or a metal foil (such as aluminumfoil), which can be anodized. Examples of glass fiber matrices includewoven and non-woven glass fiber scrims. The adsorbent sheets can be madeby coating a slurry of suitable adsorbent component, such as zeolitecrystals with binder constituents onto the reinforcement material, suchas nonwoven fiber glass scrims, woven metal fabrics, and expandedaluminum foils. In a particular embodiment, adsorbent sheets or materialare coated onto a ceramic support.

An absorber in a RCPSA unit typically comprises an adsorbent solid phaseformed from one or more adsorbent materials and a permeable gas phasethrough which the gases to be separated flow from the inlet to theoutlet of the adsorber, the components to be removed being fixed on thesolid phase. This gas phase is called “circulating gas phase” or moresimply “gas phase”. The solid phase includes a network of pores, themean size of which is usually between approximately 0.02 μm and 20 μm.There may be a network of even smaller pores, called “micropores”, thisbeing encountered, for example, in microporous carbon adsorbents orzeolites. As previously mentioned, the solid phase may be deposited on anon-adsorbent support, the function of which is to provide mechanicalstrength or support, or else to play a thermal conduction role or tostore heat. The phenomenon of adsorption comprises two main steps,namely passage of the adsorbate from the circulating gas phase onto thesurface of the solid phase, followed by passage of the adsorbate fromthe surface to the volume of the solid phase into the adsorption sites.

In an embodiment, RCPSA utilizes a structured adsorbent which isincorporated into tubes utilized in the RSPCA apparatus. Thesestructured adsorbents have an unexpectedly high mass transfer rate sincethe gas flow is through the channels formed by the structured sheets ofthe adsorbent which offers a significant improvement in mass transfer ascompared to a traditional packed fixed bed arrangement as utilized inconventional PSA. The ratio of the transfer rate of the gas phase(τ_(g)) and the mass transfer rate of the solid phase (τ_(s)) in thecurrent invention is greater than 10, preferably greater than 25, morepreferably greater than 50. These extraordinarily high mass transferrate ratios allow RCPSA to produce high purity hydrogen at a highrecovery rate with only a fraction of the equipment size, adsorbentvolume, and cost of conventional PSA.

The structured adsorbent embodiments also results in significantlygreater pressure drops to be achieved through the adsorbent thanconventional PSA without the detrimental effects associated withparticulate bed technology. The adsorbent beds can be designed withadsorbent bed unit length pressure drops of greater than 5 inches ofwater per foot of bed length, more preferably greater than 10 in.H₂0/ft, and even more preferably greater than 20 in. H₂0/ft. This is incontrast with conventional PSA units where the adsorbent bed unit lengthpressure drops are generally limited to below about 5 in. H₂0/ftdepending upon the adsorbent used, with most conventional PSA unitsbeing designed with a pressure drop of about 1 in. H₂0/ft or less tominimize the problems discussed that are associated with the largerbeds, long cycle time, and particulate absorbents of conventional PSAunits. The adsorbent beds of conventional PSA cannot accommodate higherpressure drops because of the risk of fluidizing the beds which resultsin excessive attrition and premature unit shutdowns due to accompanyingequipment problems and/or a need to add or replace lost adsorbentmaterials. These markedly higher adsorbent bed unit length pressuredrops allow RCPSA adsorbent beds to be significantly more compact,shorter, and efficient than conventional PSA.

The achievement and accommodation of the high unit length pressure dropsof the current embodiment allow high vapor velocities to be achievedacross the structured adsorbent beds. This results in a greater masscontact rate between the process fluids and the adsorbent materials in aunit of time than can be achieved by conventional PSA. This results inshorter bed lengths, higher gas phase transfer rates (τ_(g)) andimproved hydrogen recovery. With these significantly shorter bedlengths, total pressure drops of the RSCPA application of the presentinvention can be maintained at total bed pressure differentials duringthe feed cycle of about 10 to 50 psig, preferably less than 30 psig,while minimizing the active adsorbent beds to less than 5 feet inlength, preferably less than 2 feet in length and as short as less than1 foot in length.

The absolute pressure levels employed during the RCPSA process are notcritical provided that the pressure differential between the adsorptionand desorption steps is sufficient to cause a change in the adsorbatefraction loading on the adsorbent thereby providing a delta loadingeffective for separating the stream components processed by the RCPSAunit. Typical pressure levels range of the from about 50 to 2000 psia,more preferably from about 80 to 500 psia during the adsorption step.However, it should be noted that the actual pressures utilized duringthe feed, depressurization, purge and repressurization stages is highlydependent upon many factors including, but not limited to, the actualoperating pressure and temperature of the overall stream to beseparated, stream composition, and desired recovery percentage andpurity of the RCPSA product stream. U.S. Pat. Nos. 6,406,523; 6,451,095;6,488,747; 6,533,846 and 6,565,635, all of which are incorporated hereinby reference, disclose various aspects of RCPSA technology.

In an embodiment, the rapid cycle pressure swing adsorption system has atotal cycle time, t_(TOT), to separate a feed gas into product gas (inthis case, a hydrogen-enriched stream) and a tail (exhaust) gas. Themethod generally includes the steps of conducting the feed gas having ahydrogen purity F %, where F is the percentage of the feed gas which isthe weakly-adsorbable (hydrogen) component, into an adsorbent bed thatselectively adsorbs the tail gas and passes the hydrogen product gas outof the bed, for time, t_(F), wherein the hydrogen product gas has apurity of P % and a rate of recovery of R %. Recovery R % is the ratioof amount of hydrogen retained in the product to the amount of hydrogenavailable in the feed. Then the bed is co-currently depressurized for atime, t_(CO), followed by counter-currently depressurizing the bed for atime, t_(CN), wherein desorbate (tail gas or exhaust gas) is releasedfrom the bed at a pressure greater than or equal to 30 psig. The bed ispurged for a time, t_(P), typically with a portion of the hydrogenproduct gas. Subsequently the bed is repressurized for a time, t_(RP),typically with a portion of hydrogen product gas or feed gas, whereinthe cycle time, t_(TOT), is equal to the sum of the individual cycletimes comprising the total cycle time, i.e.

t _(TOT) =t _(F) +t _(CO) +t _(CN) +t _(P) +t _(RP)  (3)

This embodiment encompasses, but is not limited to, RCPSA processes suchthat either the rate of recovery, R %>80% for a product purity to feedpurity ratio, P %/F %>1.1, and/or the rate of recovery, R %>90% for aproduct purity to feed purity ratio, 0<P %/F %<1.1. Results supportingthese high recovery & purity ranges can be found in Examples 4 through10 below. Other embodiments will include applications of RCPSA inprocesses where recovery rates are much lower than 80%. Embodiments ofRCPSA are not limited to exceeding any specific recovery rate or puritythresholds and can be as applied at recovery rates and/or purities aslow as desired or economically justifiable for a particular application.

It should also be noted that it is within the scope of this inventionthat steps t_(CO), t_(CN), or t_(P) of equation (3) above can be omittedtogether or in any individual combination. However it is preferred thatall steps in the above equation (3) be performed or that only one ofsteps t_(CO) or t_(CN) be omitted from the total cycle.

In an embodiment, the tail gas is also preferably released at a pressurehigh enough so that the tail gas may be fed to another device absenttail gas compression. More preferably the tail gas pressure is greaterthan or equal to 60 psig. In a most preferred embodiment, the tail gaspressure is greater than or equal to 80 psig. At higher pressures, thetail gas can be conducted to a fuel header or directly to anotherprocess unit in a refinery or petrochemical, such as a hydroprocessingunit, a reforming unit, a fluidized catalytic cracker unit or a methanesynthesis unit. It is also within the scope of this invention for thisparticular embodiment that the only step in depressuring the bed isco-current flow. That is, the counter-current depressurizing step isomitted.

Practice of the present invention can have the following benefits:

(a) Increasing the purity of hydrogen-containing stream(s) available asmakeup gas, or of streams which must be upgraded to higher purity beforethey are suitable as make-up gas.

(b) Increasing the purity of hydrogen-containing recycle gas streamsresulting in an increase in overall hydrogen treat gas purity in thereactor to allow for higher hydrotreating severity or additional producttreating.

(c) Use for H₂ recovery from hydroprocessing purge gases, either wheresignificant concentrations of H₂S are present (before gas scrubbing) orafter gas scrubbing (typically <100 vppm H₂S).

In hydroprocessing, increased H₂ purity translates to higher H₂ partialpressures in the hydroprocessing reactor(s). This both increases thereaction kinetics and decreases the rate of catalyst deactivation. Thebenefits of higher H₂ partial pressures can be exploited in a variety ofways, such as: operating at lower reactor temperature, which reducesenergy costs, decreases catalyst deactivation, and extends catalystlife; increasing unit feed rate; processing more sour (higher sulfur)feedstocks; processing higher concentrations of cracked feedstocks;improved product color, particularly near end of run; debottleneckingexisting compressors and/or treat gas circuits (increased scf H₂ atconstant total flow, or same scf H₂ at lower total flow); and othermeans that would be apparent to one skilled in the art.

Increased H₂ recovery also offers significant potential benefits, someof which are described as follows:

(i) reducing the demand for purchased, manufactured, or other sources ofH₂ within the refinery;

(ii) increasing hydroprocessing feed rates at constant (existing) makeupgas demands as a result of the increased hydrogen recovery;

(iii) improving the hydrogen purity in hydroprocessing for increasedheteroatom removal efficiencies;

(iv) removing a portion of the H₂ from refinery fuel gas which isdetrimental to the fuel gas due to hydrogen's low BTU value which canpresent combustion capacity limitations and difficulties for somefurnace burners;

(v) Other benefits that would be apparent to one knowledgeable in theart.

Depending on the specific RCPSA design, other contaminants, such as, butnot limited to CO₂, water, and ammonia may also be removed from a feed.A portion of the scrubbed vapor stream may bypass the RCPSA unit.

The following examples are presented for illustrative purposes only andshould not be cited as being limiting in any way.

EXAMPLES Example 1

In this example, the refinery stream is at 480 psig with tail gas at 65psig whereby the pressure swing is 6.18. The feed composition andpressures are typical of refinery processing units such as those foundin hydroprocessing or hydrotreating applications. In this exampletypical hydrocarbons are described by their carbon number i.e.C₁=methane, C₂=ethane etc. The RCPSA is capable of producing hydrogenat >99% purity and >81% recovery over a range of flow rates. Tables 1aand 1b show the results of computer simulation of the RCPSA and theinput and output percentages of the different components for thisexample. Tables 1a and 1b also show how the hydrogen purity decreases asrecovery is increased from 89.7% to 91.7% for a 6 MMSCFD stream at 480psig and tail gas at 65 psig.

Tables 1a and 1b

Composition (mol %) of input and output from RCPSA (67 ft³) in H2purification. Feed is at 480 psig, 122 deg F. and Tail gas at 65 psig.Feed rate is about 6 MMSCFD.

TABLE 1a Higher purity Step Times in seconds are t_(F) = 1, t_(CO) =0.167, t_(CN) = 0, t_(P) = 0.333, t_(RP) = 0.5 H2 at 98.6% purity, 89.7%recovery FEED PRODUCT TAIL-GAS H2 88.0 98.69 45.8 C1 6.3 1.28 25.1 C20.2 0.01 1.0 C3 2.6 0.01 12.3 C4+ 2.9 0.00 14.8 H2O 2000 vppm 65 vppm9965 vppm TOTAL 6.162 4.934 1.228 (MMSCFD) 480 psig 470 psig 65 psig

TABLE 1b Higher purity Step Times in seconds are t_(F) = 1, t_(CO) =0.333, t_(CN) = 0, t_(P) = 0.167, t_(RP) = 0.5 H2 at 97.8% purity, 91.7%recovery FEED PRODUCT TAIL-GAS H2 88.0 97.8 45.9 C1 6.3 2.14 25.0 C2 0.20.02 1.0 C3 2.6 0.02 12.3 C4+ 2.9 0.00 14.9 H2O 2000 vppm 131 vppm 10016vpm TOTAL 6.160 5.085 1.074 (MMSCFD) 480 psig 470 psig 65 psig

The RCPSA's described in the present invention operate a cycleconsisting of different steps. Step 1 is feed during which product isproduced, step 2 is co-current depressurization, step 3 iscounter-current depressurization, step 4 is purge, usuallycounter-current) and step 5 is repressurization with product. In theRCPSA's described here at any instant half the total number of beds areon the feed step. In this example, t_(TOT)=2 sec in which the feed time,t_(F), is one-half of the total cycle.

Example 2

In this example, the conditions are the same as in Example 1. Table 2ashows conditions utilizing both a co-current and counter-current stepsto achieve hydrogen purity >99%. Table 2b shows that the counter-currentdepressurization step may be eliminated, and a hydrogen purity of 99%can still be maintained. In fact, this shows that by increasing the timeof the purge cycle, t_(P), by the duration removed from thecounter-current depressurization step, t_(CN), that hydrogen recoverycan be increased to a level of 88%.

Tables 2a and 2b

Effect of step durations on H2 purity and recovery from an RCPSA (67ft³). Same conditions as Table 1. Feed is at 480 psig, 122 deg F. andTail gas at 65 psig. Feed rate is about 6 MMSCFD.

TABLE 2a With counter-current depress, Intermediate pressure = 105 psig.Purity Recovery t_(F) t_(CO) t_(CN) t_(P) t_(RP) % % s S s S S 98.2 84.31 0.283 0.05 0.167 0.5 98.3 85 1 0.166 0.167 0.167 0.5 99.9 80 1 0.0830.25 0.167 0.5

TABLE 2b Without counter-current depress Purity Recovery t_(F) t_(CO)t_(CN) t_(P) t_(RP) % % s S s S s 97.8 91.7 1 0.333 0 0.167 0.5 98.7 901 0.166 0 0.334 0.5 99 88 1 0.083 0 0.417 0.5

Example 3

This example shows a 10 MMSCFD refinery stream, once again containingtypical components, as shown in feed column of Table 3 (e.g. the feedcomposition contains 74% H₂). The stream is at 480 psig with RCPSA tailgas at 65 psig whereby the absolute pressure swing is 6.18. Once againthe RCPSA of the present invention is capable of producing hydrogenat >99% purity and >85% recovery from these feed compositions. Tables 3aand 3b show the results of this example.

Tables 3a and 3b

Composition (mol %) of input and output from RCPSA (53 ft³) in H2purification. Feed is at 480 psig, 101 deg F. and Tail gas at 65 psig.Feed rate is about 10 MMSCFD.

TABLE 3a Higher purity Step Times in seconds are t_(F) = 0.583, t_(CO) =0.083, t_(CN) = 0, t_(P) = 0.25, t_(RP) = 0.25 H2 at 99.98% purity and86% recovery FEED PRODUCT TAIL-GAS H2 74.0 99.98 29.8 C1 14.3 0.02 37.6C2 5.2 0.00 13.8 C3 2.6 0.00 7.4 C4+ 3.9 0.00 11.0 H2O 2000 vppm 0.3vppm 5387 vppm TOTAL 10.220 6.514 3.705 (MMSCFD) 480 psig 470 psig 65psig

TABLE 3b Lower purity Step Times in seconds are t_(F) = 0.5, t_(CO) =0.167, t_(CN) = 0, t_(P) = 0.083, t_(RP) = 0.25 H2 at 93% purity and 89%recovery FEED PRODUCT TAIL-GAS H2 74.0 93.12 29.3 C1 14.3 6.34 31.0 C25.2 0.50 16.6 C3 2.6 0.02 8.9 C4+ 3.9 0.00 13.4 H2O 2000 vppm 142 vppm6501 vppm TOTAL 10.220 7.240 2.977 (MMSCFD) 480 psig 470 psig 65 psig

In both cases shown in Tables 3a and 3b above, although tail gaspressure is high at 65 psig, the present invention shows that highpurity (99%) may be obtained if the purge step, t_(P), is sufficientlyincreased.

Tables 2a, 2b and 3a show that for both 6 MMSCFD and 10 MMSCFD flow rateconditions, very high purity hydrogen at ˜99% and >85% recovery isachievable with the RCPSA. In both cases the tail gas is at 65 psig.Such high purities and recoveries of product gas achieved using theRCPSA with all the exhaust produced at high pressure have not beendiscovered before and are a key feature of the present invention.

Table 3c shows the results for an RCPSA (volume=49 cubic ft) thatdelivers high purity (>99%) H₂ at high recovery for the same refinerystream discussed in Tables 3a and 3b. As compared to Table 3a, Table 3cshows that similar purity and recovery rates can be achieved bysimultaneously decreasing the duration of the feed cycle, t_(F), and thepurge cycle, t_(P).

TABLE 3c Effect of step durations on H2 purity and recovery from anRCPSA (49 ft³). Feed is at 480 psig, 101 deg F. and Tail gas at 65 psig.Feed rate is about 10 MMSCFD. Without counter-current depress. PurityRecovery t_(F) t_(CO) t_(CN) t_(P) t_(RP) % % s S S s s 95.6 87.7 0.50.167 0 0.083 0.25 97.6 86 0.5 0.117 0 0.133 0.25 99.7 85.9 0.5 0.083 00.167 0.25

Example 4

In this example, Table 4 further illustrates the performance of RCPSA'soperated in accordance with the invention being described here. In thisexample, the feed is a typical refinery stream and is at a pressure of300 psig. The RCPSA of the present invention is able to produce 99% purehydrogen product at 83.6% recovery when all the tail gas is exhausted at40 psig. In this case the tail gas can be sent to a flash drum or otherseparator or other downstream refinery equipment without furthercompression requirement. Another important aspect of this invention isthat the RCPSA also removes CO to <2 vppm, which is extremely desirablefor refinery units that use the product hydrogen enriched stream. Lowerlevels of CO ensure that the catalysts in the downstream units operatewithout deterioration in activity over extended lengths. ConventionalPSA cannot meet this CO specification and simultaneously also meet thecondition of exhausting all the tail gas at the higher pressure, such asat typical fuel header pressure or the high pressure of other equipmentthat processes such RCPSA exhaust. Since all the tail gas is availableat 40 psig or greater, no additional compression is required forintegrating the RCPSA with refinery equipment.

TABLE 4 Composition (mol %) of input and output from RCPSA (4 ft³) incarbon monoxide and hydrocarbon removal from hydrogen. Feed is at 300psig, 101 deg F., and Feed rate is about 0.97 MMSCFD. Step Times inseconds are t_(F) = 0.5, t_(CO) = 0.1, t_(CN) = 0, t_(P) = 0.033, t_(RP)= 0.066 H2 at 99.99% purity and 88% recovery FEED PRODUCT TAIL-GAS H289.2 99.98 48.8 C1 3.3 0.01 13.9 C2 2.8 0.01 13.9 C3 2.0 0.00 10.2 C4+2.6 0.00 13.2 CO 50 1.1 198.4 TOTAL 0.971 0.760 0.211 300 psig 290 psig40 psig

Example 5

Tables 5a and 5b compare the performance of RCPSA's operated inaccordance with the invention being described here. The stream beingpurified has lower H₂ in the feed (51% mol) and is a typicalrefinery/petrochemical stream. In both cases (corresponding to Tables 5aand 5b), a counter current depressurization step is applied after theco-current step. In accordance with the invention, Table 5a shows thathigh H₂ recovery (81%) is possible even when all the tail gas isreleased at 65 psig or greater. In contrast, the RCPSA where sometail-gas is available as low as 5 psig, loses hydrogen in thecounter-current depressurization such that H₂ recovery drops to 56%. Inaddition, the higher pressure of the stream in Table 5a indicates thatno tail gas compression is required.

Tables 5a and 5b

Effect of Tail Gas Pressure on recovery. Example of RCPSA applied to afeed with H2 concentration (51.3 mol %). Composition (mol %) of inputand output from RCPSA (31 ft³) in H2 purification. Feed is at 273 psig,122 deg F. and Feed rate is about 5.1 MMSCFD.

TABLE 5a Step Times in seconds are t_(F) = 0.5, t_(CO) = 0.083, t_(CN) =0.033, t_(P) = 0.25, t_(RP) = 0.133 [A] Tail gas available from 65-83psig, H2 at 99.7% purity and 81% recovery FEED PRODUCT TAIL-GAS H2 51.399.71 20.1 C1 38.0 0.29 61.0 C2 4.8 0.00 8.0 C3 2.2 0.00 3.8 C4+ 3.70.00 6.4 H20 4000 vppm 0.7 vppm 6643 vppm TOTAL 5.142 2.141 3.001(MMSCFD) 273 psig 263 psig 65-83 psig

TABLE 5b Step Times in sec. are t_(F) = 0.667, t_(CO) = 0.167, t_(CN) =0.083, t_(P) = 0.083, t_(RP) = 0.33 [B] Tail gas available from 5-65psig, H2 at 99.9% purity and 56% recovery FEED PRODUCT TAIL-GAS H2 51.399.99 34.2 C1 38.0 0.01 48.8 C2 4.8 0.00 6.9 C3 2.2 0.00 3.4 C4+ 3.70.00 6.2 H20 4000 vppm 0.0 vppm 5630 vppm TOTAL 5.142 1.490 3.651(MMSCFD) 273 psig 263 psig 5-65 psig

Example 6

In this example, Tables 6a and 6b compare the performance of RCPSA'soperated in accordance with the invention being described here. In thesecases, the feed pressure is 800 psig and tail gas is exhausted at either65 psig or at 100 psig. The composition reflects typical impurities suchH2S, which can be present in such refinery applications. As can be seen,high recovery (>80%) is observed in both cases with the highpurity >99%. In both these cases, only a co-current depressurization isused and the effluent during this step is sent to other beds in thecycle. Tail gas only issues during the countercurrent purge step. Table6c shows the case for an RCPSA operated where some of the tail gas isalso exhausted in a countercurrent depressurization step following aco-current depressurization. The effluent of the co-currentdepressurization is of sufficient purity and pressure to be able toreturn it one of the other beds in the RCPSA vessel configuration thatis part of this invention. Tail gas i.e., exhaust gas, issues during thecounter-current depressurization and the counter-current purge steps.

In all cases the entire amount of tail gas is available at elevatedpressure which allows for integration with other high pressure refineryprocess. This removes the need for any form of required compressionwhile producing high purity gas at high recoveries. In accordance withthe broad claims of this invention, these cases are only to beconsidered as illustrative examples and not limiting either to therefinery, petrochemical or processing location or even to the nature ofthe particular molecules being separated.

Tables 6a, 6b, and 6c

Example of RCPSA applied to a high pressure feed. Composition (mol %) ofinput and output from RCPSA (18 ft³) in H2 purification. Feed is at 800psig, 122 deg F. and Feed rate is about 10.1 MMSCFD.

TABLE 6a Step Times in seconds are t_(F) = 0.91, t_(CO) = 0.25, t_(CN) =0, t_(P) = 0.33, t_(RP) = 0.33 [A] Tail gas at 65 psig, H2 at 99.9%purity and 87% recovery FEED PRODUCT TAIL-GAS H2 74.0 99.99 29.5 C1 14.30.01 37.6 C2 5.2 0.00 14.0 C3 2.6 0.00 7.4 C4+ 3.9 0.00 10.9 H20 20 vppm0 55 vppm TOTAL 10.187 6.524 3.663 (MMSCFD) 800 psig 790 psig 65 psig

TABLE 6b Step Times in seconds are t_(F) = 0.91, t_(CO) = 0.25, t_(CN) =0, t_(P) = 0.33, t_(RP) = 0.33 [B] Tail gas at 100 psig, H2 at 99.93%purity and 80.3% recovery FEED PRODUCT TAIL-GAS H2 74.0 99.93 38.1 C114.3 0.07 32.8 C2 5.2 0.00 12.5 C3 2.6 0.00 6.5 C4+ 3.9 0.00 9.6 H2S 20vppm 0 vppm 49 vppm TOTAL 10.187 6.062 4.125 (MMSCFD) 800 psig 790 psig100 psig

TABLE 6c Step times in seconds are t_(F) = 0.91, t_(CO) = 0.083, t_(CN)= 0.25, t_(P) = 0.167, t_(RP) = 0.41 [C] Tail gas at 65-100 psig, H2 at99.8% purity and 84% recovery FEED PRODUCT TAIL-GAS H2 74.0 99.95 28.9C1 14.3 0.05 39.0 C2 5.2 0.00 13.7 C3 2.6 0.00 7.2 C4+ 3.9 0.00 10.6 H2S20 vppm 0.01 vppm 53 vppm TOTAL 10.187 6.373 3.814 (MMSCFD) 800 psig 790psig 65-100 psig

Example 7

Tables 7a, 7b, and 7c compare the performance of RCPSA's operated inaccordance with the invention being described here. The stream beingpurified has higher H₂ in the feed (85% mol) and is a typicalrefinery/petrochemical stream. In these examples the purity increase inproduct is below 10% (i.e. P/F<1.1). Under this constraint, the methodof the present invention is able to produce hydrogen at >90% recoverywithout the need for tail gas compression.

Tables 7a, 7b, and 7c

Example of RCPSA applied to a Feed with H2 concentration (85 mol %).Composition (mol %) of input and output from RCPSA (6.1 ft³). Feed is at480 psig, 135 deg F. and Feed rate is about 6 MMSCFD.

TABLE 7a Step Times in seconds are t_(F) = 0.5, t_(CO) = 0.33, t_(CN) =0.167, t_(P) = 0.167, t_(RP) = 1.83 recovery = 85% FEED PRODUCT TAIL-GASH2 85.0 92.40 57.9 C1 8.0 4.56 17.9 C2 4.0 1.79 13.1 C3 3.0 1.16 10.4C4+ 0.0 0.00 0.0 H2O 2000 866.5 6915 TOTAL 6.100 4.780 1.320 (MMSCFD)480 psig 470 psig 65 psig

TABLE 7b Step Times in sec. are t_(F) = 1, t_(CO) = 0.333, t_(CN) =0.167, t_(P) = 0.083, t_(RP) = 0.417 recovery = 90% FEED PRODUCTTAIL-GAS H2 85.0 90.90 58.2 C1 8.0 5.47 18.1 C2 4.0 2.23 12.9 C3 3.01.29 10.1 C4+ 0.0 0.00 0.0 H2O 2000 1070.5 6823 TOTAL 6.120 5.150 0.969(MMSCFD) 480 psig 470 psig 65 psig

TABLE 7c Step Times in sec. are t_(F) = 2, t_(CO) = 0.667, t_(CN) =0.333, t_(P) = 0.167, t_(RP) = 0.833 recovery = 90% FEED PRODUCTTAIL-GAS H2 85.0 90.19 55.2 C1 8.0 6.21 18.8 C2 4.0 2.32 13.9 C3 3.01.17 11.3 C4+ 0.0 0.00 0.0 H2O 2000 1103.5 7447 TOTAL 6.138 5.208 0.93(MMSCFD) 480 psig 470 psig 65 psig

1. A process for upgrading a hydrocarbon feed in a hydroprocessingprocess unit, comprising: a) contacting said hydrocarbon feed in ahydroprocessing zone with hydrogen, a portion of which is obtained froma hydrogen-containing make-up gas, and a catalytically effective amountof a hydroprocessing catalyst at hydroprocessing conditions therebyresulting in a liquid phase and a vapor phase; b separating said liquidphase product and said vapor phase, which vapor phase product containshydrogen and light hydrocarbons; c) removing at least a portion of thelight hydrocarbons from the hydrogen-containing make-up treat gas, thevapor phase product, or both, in a rapid cycle pressure swing adsorptionunit containing a plurality of adsorbent beds and having a total cycletime of less than about 30 seconds and a pressure drop within eachadsorbent bed of greater than about 5 inches of water per foot of bedlength; d) recycling at least a portion of the vapor phase of step c)above having a higher concentration of hydrogen to the hydroprocessingzone.
 2. The process of claim 1 wherein the hydrocarbon feed is selectedfrom the group consisting of naphtha boiling range feeds, kerosene andjet fuel boiling range feeds, distillate boiling range feeds, residesand crudes.
 3. The process of claim 2 wherein the hydrocarbon feed is anaphtha boiling range feed selected from the group consisting ofstraight run naphtha, cat cracked naphtha, coker naphtha, hydrocrackernaphtha, resid hydrotreater naphtha.
 4. The process of claim 2 whereinthe hydrocarbon feed is a distillate and higher boiling range feedselected from the group consisting of cycle oils produced from the FluidCatalytic Cracker (FCC), atmospheric and vacuum gas oils, atmosphericand vacuum residua, pyrolysis gasoline, Fischer-Tropsch liquids,raffinates, waxes, lube oils, and crudes.
 5. The process of claim 1wherein the hydroprocessing processing is a hydrotreating process forthe removal of heteroatoms from the hydrocarbon feed and whereinhydrogen sulfide is also a component of the vapor phase and wherein atleast a portion of the hydrogen sulfide is removed in an acid gasscrubbing zone prior to having light hydrocarbons removed by rapid cyclepressure swing adsorption.
 6. The process of claim 5 wherein the totalcycle time or rapid cycle pressure swing adsorption is less than about15 seconds.
 7. The process of claim 6 wherein the total cycle time isless than about 10 seconds and the pressure drop of each adsorbent bedis greater than about 10 inches of water per foot of bed length.
 8. Theprocess of claim 7 wherein the total cycle time is less than about 5seconds.
 9. The process of claim 8 wherein the pressure drop of greaterthan about 20 inches of water per foot of bed length.
 10. The process ofclaim 1 wherein the cycle time is less than about 10 seconds and thepressure drop is greater than about 10 inches of water per foot of bedlength.
 11. The process of claim 10 wherein the cycle time is less thanabout 5 seconds and the pressure drop is greater than about 20 inches ofwater per foot of bed length.
 12. The process of claim 1 wherein thehydroprocessing process is hydrocracking wherein a hydrocarbon feed isconverted to lower boiling products.
 13. The process of claim 12 whereinthe total cycle time or rapid cycle pressure swing adsorption is lessthan about 15 seconds.
 14. The process of claim 13 wherein the totalcycle time is less than about 10 seconds and the pressure drop of eachadsorbent bed is greater than about 10 inches of water per foot of bedlength.
 15. The process of claim 14 wherein the total cycle time is lessthan about 5 seconds.
 16. The process of claim 15 wherein the pressuredrop of greater than about 20 inches of water per foot of bed length.17. The process of claim 1 wherein the hydroprocessing process ishydroisomerization wherein molecules of the hydrocarbon feed areisomerized.
 18. The process of claim 17 wherein the total cycle time orrapid pressure swing adsorption is less than about 15 seconds.
 19. Theprocess of claim 18 wherein the total cycle time is less than about 10seconds and the pressure drop of each adsorbent bed is greater thanabout 10 inches of water per foot of bed length.
 20. The process ofclaim 19 wherein the total cycle time is less than about 5 seconds. 21.The process of claim 20 wherein the pressure drop of greater than about20 inches of water per foot of bed length.
 22. The process of claim 1wherein the hydroprocessing process is hydrogenation wherein unsaturatesand aromatics are hydrogenated.
 23. The process of claim 22 wherein thetotal cycle time or rapid pressure swing adsorption is less than about15 seconds.
 24. The process of claim 23 wherein the total cycle time isless than about 10 seconds and the pressure drop of each adsorbent bedis greater than about 10 inches of water per foot of bed length.
 25. Theprocess of claim 24 wherein the total cycle time is less than about 5seconds.
 26. The process of claim 24 wherein the pressure drop ofgreater than about 20 inches of water per foot of bed length.